Pretreatment and cracking of residual oils



Dec. 22, 1964 A. D. ANDERSON ETAL 3,162,596 PRETREATMENT AND cRAcKING oFREsIuUAL oILs Filed July 24. 1961 i D: 2 J3 `I as g .v L 1J L r 2 g g an: el N LAJ LLI @1 o q' O sa o "1 w w e, v 3 N m\ O f2 f *l E f r :N 5 Lt o f\ l v 2 \ic"\. U3 o S ,E NE

' Ql-H t :cu2J 9 s Q u 9m INVENToRs L ARvlN D` ANDERSON ROBERT A.SANFORD NE YS.

United States Patent O il s.

3,i6.,596 PRETREATMENT AND CRACKNG F RESEUAL llLS Arvin D. Anderson,Santa Ana, falli., and Robert A.

Sanford, Homewood, lill., assignors, by mesne assignments, to SinclairResearch, lne., New York, NX., a

corporation of Eeiawme Filed .l'uly 24, 196i, Ser. No. 6,331 14 iaims.(Cl. 20d-8%) This invention relates to the upgrading of hydrocarbonoils, particularly high boiling petroleum oils, to obtain lightercomponents including gasoline of relatively high octane number. Morespecii'ically, this invention cornprises an integrated process forhydrogenating and vacnum distilling residual oils to obtain ieedstockscapable or being catalytically converted to lighter boiling material.The cracking system is coupled with a catalyst demetallization system.

The catalytic cracking or various heavier mineral hydrocarbons, forinstance, petroleum or other mineral oil distillates such as straightrun and cracked gas oils; shale oils; petroleum residues, etc., has beenproposed for many years and the catalytic cracking of gas oils ispracticed commercially to a considerable extent. As is well known tothose familiar with the art, gas oil is a broad, general term thatcovers a variety of stocks. The term includes an raction distilled frompetroleum or other mineral oil which has an initial boiling point of atleast about 400 F., and an end boiling point of at least about 600 F.,and boiling over a range of at least about 100 F. The portion which isnot distilled before the end point is reached is considered residualstock. The exact boiling range of a gas oil, therefore, will bedetermined by the initial distillation temperature (initial boilingpoint) and by the temperature at which distillation is cut off (endboiling point). ln practice, petroleum distillations have been madeunder vacuum up to temperatures as high as about MOG-1200" F. (correctedto atmospheric pressure). Accordingly, in the broad sense, a gas oil isa petroleum fraction which boils substantially continuously between twotemperatures that establish a range falling within from about 400 F. toabout ll00-l200 F. Thus, a gas oil could boil over the entire rangeabout 400-1 200 F. or it could boil over a narrower range, eg., about 5(l0-900 F.

The gas oils can be further roughly classified by boiling ranges. Thus,gas oil boiling between about m0-500 F. and about 60C-650 F. is termed alight gas oil; a medium gas oil distills between about 600650 F. andabout 800- 900" a gas oil boiling between about 80G-S50 F. and about110D-1200" F. is sometimes designated as a vacuum gas oil because itmust be distilled under severelyr reduced pressures to avoid cracking orother high temperature effects. lt must be understood, however, that aparticular stock may bridge two boiling ranges, or even span severalranges, i.e., include, for example, light and medium gas oils.

A residual stock is in general any petroleum fraction higher boilingthan gas oils which are undistilled. Any fraction, regardless oi itsinitial boiling point, which includes the heavy bottoms, such as tars,asphalte, or other undistillated materials may be termed a residualfraction. Accordingly, a residual stock can be the portion or" the cruderemaining undistilled at about ll00-l200 F., or it can be made up of agas oil fraction plus the portion undistilled at about 11001-1200" F.For instance, a topped crude may be the entire portion of the cruderemaining after the light ends (the portion boiling up to about 400 F.)have been removed by distillation. Therefore, such a fraction includesthe entire gas oil fraction (400 F. to 11004200" F.) and the undistilledportion of the crude etroleum boiling above HOO-1200 F.

The behavior of a hydrocarbon feedstock in the craclo ing reactionsdepends upon various factors including its boiling point, carbon-formingtendencies, content of cata# lyst contaminating metal-s, etc., and thesecharacteristics may aiiect the operation to an extent which maires `agiven feedstock uneconornical to employ. By and large, residual stockshave not been catalytically cracked on a commercial scale as theircarbon-forming tendencies and catalyst poisoning metals content aregenerally too great. Moreover, even some distillate materials needimprovement in their hydrogen-to-carbon ratio or contain excessiveamounts of metals which mitigates their usefulness in cracking.Frequently the rel'iner may take specialv cuts of a crude oil whichcontains metal poisons and pretreat them prior to cracking in order thatthe cracking operation itself becon es more trouble-free overall eventhough pretreatment adds to the cost and may sometimes reduce the amountof cracking feed per barrel of crude oil and thereby diminish the yieldof gasoline or be ineiiicie'nt in removing metal poisons, especiallyfrom a residual. Although the cracking catalyst employed can bediscarded citen to prevent a high accumulation of poisoning metals inthe cracking system, this type of operation represents a substantialcost actor. improvements in the feedstock characteristics become evenmore important as the cost of the catalyst rises and thus the eiects oflow feedstock quality are particularly burdensome in systems` employing'cracking catalysts containing relatively expensive syntheticcomponents.

Attempts to employ heavier fractions of crude oil for catalytic crackinghave been limited heretofore due to the presence of metal contaminantsin such heavy fractions. The highest boiling fractions of a crude oilcontain substantial portions of metal contaminants, Vparticularly nickeland vanadium. When an attempt is made to segregate higher boilingdistillate fractions of a crude oil, some portion of these metalcontaminants and colte-formers is inherently and unavoidably carriedover into the distillate product. For example, when a heavy gas oildistillate fraction, having a boiling range of about 700 to i E., issegregated from a crude oil, about 0.1 to l pound per 1000 barrels ofmetal contaminants sometimes can be'obtained from the gas oil distillatein a typical distillation.

The problem of metal ycontaminant carryover in the segregation of heavydistillate fractions is believed to be due to two phenomena. First ofall, it is believed that the metal contaminants occur or are convertedduring `distillation to the form oi metal complexes. These cornplexesmay generally be identified as large condensed ring substances. Some ofthese metal complexes and particularly nickel porphyrins are sudicientlyvolatile so as to be carried overhead under vacuum distillationconditions. Consequently, when attempting to segregate heavy boil inggas oil fractions including components boiling above about 700 or 900F., volatile metal contaminants are unavoidably obtained in thedistillate product. The second phenomenon is generally referred to asmechanical entrainment. In distillation a small portion of high boilingliquid hydrocarbons from the residual fraction is normally entrained inthe lower-boiling componentsv taken overhead in a distillationoperation. Since liquid hydrocarbons derived from residuals containconcentrated amounts of metal contaminants, such entrainmentindistillate products accounts for a portion of the metal contamiriationof such distillates.

Since essentially fewer metals are found in most petroleum distillateswhich boil below about 700 or 900 F., petroleum refiners are frequentlylimited to cracking stocks in this lower boiling range recoverable froma crude source by distillation at atmospheric or slightly reducedpressures. The metal poisons and coke-orrners are for the most part leftbehind in the non-distilled, residue portion of the oil. Refiners,therefore, lare usually provided with a distillate fraction of the crudewhich is relativelylow in metals and coke formers to serve as feed tocatalytic cracking. However, ythere is available in most crudes aconsiderable amount of oil, boiling in the range of about 700 to 1200 F.or higher, which would be a good Cracking stock if it could be recoveredsufficiently free of metal poisons land coke formers. Yet Vacuumdistillation of the highly metals-contaminated residualfrequentlycarries a good deal of metals over into the heavy gas oil distillatethereby obtained. In this invention the metals con-V tent of thefeedstock is reduced by hydrogenation prior to the vacuum distillation,offering higher boiling components less contaminated with metals andcoke-formers which are subject to volatilization `and/o1" entrainment inthe vacuum distillation.

It has been proposed heretofore to hydrogenate the various heavymetal-containing hydrocarbon oils prior to charging them, or a fractionthereof, to a catalytic cracking operation. By so doing the hydrocarbonmay be given an improved hydrogen-to-carbon ratio and the amount ofcontaminants, such as coke-formers, sulfur, and nitrogen may be reduced.The content of metals which poison cracking catalysts is also reducedand removal of any n substantial amount of these contaminants from thecracking feed tends to enhance eiciency of the catalytic crackingoperation. The degree of feedstock improvement from hydrogenationisdependent, however, upon several factors which include the severity ofreaction and the amount of hydrogen consumed. High severity can increasethe extent of metals removal; however, Ait may involve a greater:consumption of hydrogen, -a larger capital investment for high-pressureequipment, and a reduction in the yield of cracking feedstock.

Petroleum fractions containing large amounts of cokeforming and/ ormetal components, such as the asphaltic and residual materials describedabove, frequently require such severe hydrotreating to make themtrouble-y free cracliing feeds in conventional processes that theexpense of such hydrotreating is not practical. This invention employspartial demetallization of a cracking feed by hydrogenating anddistillation along with demetallization of the cracking catalyst byprocedures to be described, to achieve greater economy than would beobtained by employing only one of the demetallizing'techniques:hydrogenation and/ or distillation on the one hand, or poison removalfrom the cracking catalyst, on the other hand, in the yattempt toobviate poisoning eifects. In the operation of this invention,hydrogenation, distillation and catalytic cracking of heavier mineral.hydrocarbon oil feedstocks to produce gasoline are combined andbalancedl with a procedure for reducing poisoning metals on the crackingcatalyst to present a much more attractive alternative to the individualopenations described abovea for overcoming an overall metals problem.Under these conditions all four of the hydrogenation, distillation,catalytic cracking and cracking catalyst demetallization can be operatedto make a relatively low consumption of hydrogen during hydrogenationmore attractive, with the hydrogen being better utilized lthroughminimization of dehydrogenation in the catalytic cracking operation andthe vacuum distillation less likely to produce a heavy gas oil toocontaminated for use in cracking. In this method the metal-containinghydrocarbon feedstock is hydrogenated under conditions giving a desiredhydrogenation elfect with partial, but not complete, removal ofpoisoning metals. The hydrogenated oil or a selected portion thereofboiling above about 400 F. and containing a significant amount of metalcontaminant is distilled to obtain a heavy gas oil fraction, containingsubstantially less contaminants than thehydrogenated product, which isthen catalytically cracked by itself or after blending with conventionalcracking fecdstocks to produce in good yield a gasoline fraction ofrelatively high octane rating.

Ak As the cracking operation proceeds the catalyst is treated to removeaccumulated metal poisons and is then reused in the cracking operation.Besides preserving the effects of hydrogenation by minimizing thedehydrogenation effects of a poisoned catalyst, catalyst demetallizationprovides a cracking operation in which there is relatively less carbonlaydown on the catalyst. This further increases gasoline yield in asystem having a given carbonburuing capacity. The cracking aspect ofthis invention With its demetallization features is economicallyattractive when a cracking feedstock is obtained containing as little asabout 0.5 ppm. nickel and/ or about 0.5 p.p.m. vanadium.V

The feeds to the present process comprise petroleum residua which Vmaybe exemplified by vacuum residua, atmospheric residua, tars, pitches,etc., boiling primarily above about 600 F. or even above about 900 F.The

, residual feed often has an API gravity in the range of about 0 to 25,aConradson carbon content in the range of about 3 to 35 weight percentand a viscosity often above about 200 seconds Saybolt Furol at 210 F.These charge stocks contain metals which are poisonous tothe K contentsabove these ranges may be present; it will be ap- Yeral crackingcharacteristics.

parent that oils having metal contents in these generally undesirableranges are the oils which this invention salvages. In most cases, thetotal of one or both of these metals in the residual will be at leastabout 5 ppm. and may often contain about 25 or 50 p.p.m. nickel andabout or 100 ppm. vanadium. The maximum amount of metals in theresiduals can vary Widely. The maximum amount of these poisoning metalsin the residual stock will usually not exceed about 500 ppm. nickel,and/ or about 1000 p.p.m.'vanadium, to be economically processed.Residuals containing a low level of metals contaminants profit by thehydrogenation and distillation step of this invention by improvements intheir gen- Although referred to =as metals the contaminants may be inthe form of free metals or metal compounds and it is to be understoodthat the term metal used herein refers to either form.

In' hydrogenation the hydrogen has a number of eects on the feedstock.One use is to dissociate the heavy poisoning metals from their compoundsin the feedstock.

Y materials.

Hydrogenation also serves to saturate components of the feed which aresusceptible to such. Thus hydrogenation increases the hydrogen-to-carbonratio of the hydrogenated eiiluent reducing the coke-forming tendenciesof the feedstock. Also, hydrogenation may cause a certain amount of thefeed to be converted (cracked) to lower boiling materials. This factormay sometimes make residual oils a more desirable feedstock for thisinvention than lower boiling materials of the same metal content. Sincemore cracking is not only permissible but indeed sometimes desirablewhen hydrogenation is performed on a residual to prepare it forcatalytic cracking, the hydrogenation may be performed at highertemperatures.

lAtthese temperatures, hydrogenationcan perform its denietallization andsaturation functions Vat a lower pressure, thereby lowering equipmentcosts. Any gas or gasoline produced in the hydrogenation and unsuitablefor use as feed to the catalytic cracking may be used as refomerfeedstock, although hydrogenation of residuals usually does not produceoverly large amounts of these Usually greater than about 10% of theresidual hydrocarbon charge may be converted to lower boiling normallyliquid materials by hydrogenation and quite frequently at least about25% is converted. Conversion to lower boiling materials rarely exceedsof a residual charge. Y

In the hydrogenation operation hydrogen is consumed by chemicalcombination With a component of the hydrocarbon feed. Wherehydrocracking is performed the hydrogen consumption will be high, as inthe case of treating most residual hydrocarbon oils. Hydrogenationoperations are subject to various modifications from which the petroleumreiiner may choose, depending upon the residual stock to be treated andthe results desired. Hydrogenation may be performed by free or molecularhydrogen in the presence of a catalyst or by a hydrogen donatingchemical with or without a catalyst. Elevated temperature (about400-l200 F.) and pressure (atmospheric to 3000 p.s.i.g.) conditionsusually prevail in hydrogenation, and Within these ranges conditions canbe chosen to give the desired degree of oil demetallization, saturationand/ or cracking.

Hydrogen donor diluent craclcing (HDDC) is widely known in the art andis illustrated in abandoned application Serial No. 365,335, filed July1, 1953, by Arthur W. Langner, Jr. as disclosed in US. Patent 2,772,718.The donor diluent is a material, generally a hydrocarbon, which has theability to take up hydrogen in a hydrogenation zone and readily releaseit in a thermal treating zone. It is believed that the donor diluentoperates by yielding hydrogen atoms to the radicals that have beencreated from the residuum by the thermal treatment, thereby upgradingthe residuum and preventing condensation and/or polymerization of theradicals. In donor diluent operations, the donor diluent material issubstantially unaltered as it passes through the process, and it isusually' customary to recycle the material so that it is used over andover again as a hydrogen carrier. Donor materials may be added as arelatively pure chemical such as tetralin or Decalin or in admixturewith ot. er materials, particularly hydrocarbons, or the donor diluentmay be a partially hydrogenated catalytic cycle oil, a partiallyhydrogenated lubricating oil extract or other partially hydrogenatedaromatic. Hydrogen donors usually contain condensed ring aromatics insuticient quantities to serve as a hydrogen carrier. These aromatics arepartially hydrogenated; there is added to them some easily removablehydrogen atoms but not enough to convert the aromatics substnatially tonaphthenes. This material after being partially hydrogenated, can beadmixed with the feedstock to this process and the mixture thermallytreated, whereby the hydrogen is transferred from the partiallyhydrogenated material to the hydrogen-deficient residuum.

After blending of hydrogen donor diluent and residuum, the blend is thentaken to a thermal or catalytic cracking zone where recycle oil from thecracker, fractionator or distillation zone of this process may alsobeadded. The mixture of residual and diluent may be thermally cracked byheating to a temperature of about 700 to 1200" F., preferably about S to1000 F. at pressures within the range of from about atmospheric to about2000 psig., preterably froni about atmospheric to about 200 p.s.i.g.with a holding time of about to 30 minutes. A conventional coil or coiland drum heater may be used. A wide range of conversion may be obtainedby varying the temperature or feed rate in the thermal crackingoperation depending upon tbe feed to be processed, to produce from about40 lto S0 percent conversion to lower boiling Imaterials per pass. Theweight ratio of diluent to residual plus recycle bottoms is usuallybetween about 0.1/1 to about 1, preferably about 0.5 to Z/l. A high rateof diluent to feed and a moderately high hydrogen content tend to reducecoke for-mation and remove metal contaminants at any severity. However,the severity of thermal cracking is primarily a function of crackingtemperature and feed rate. The nature of the residual and its priorprocessing, if any, may meer the cracking severity and the amo-unt ofmetal removal.

In the catalyzed donor cracking method (CDC) a petroleum residual can becontacted at elevated temperatures with an essentially anhydrouscatalyst comprising the hydride yof a halogen having an atomic number of35 to 53 and in the presence of a hydroaro-matic material. As a result,valuable low molecular Weight hydrocarbon gases primarily boiling belowC3 and gasoline having an end boiling point of about 430 F. are producedalong with a good yield of gas oils `which provide a high qualitycatalytic cracking stock, and a residual oil boiling primarily above 950F. which is reduced in contaminating metals. The catalyst in thisprocess is hydrogen iodide or hydrogen bromide. These can be added assuch to the reaction zone or the corresponding elemental halogen, thatis, bromine or iodine, or other material which gives the halogen hydridemay be charged. The halogen hydride `is apparently in equilibrium in thereaction zone with elemental halogen. The halogen hydride maypredominate in the equilibrium mixture. The catalyst selected isemployed in -a substantially anhydnous form although it may be used insolution with alcohol or other solvents. The amount of catalyst utilizednormally depends on the characteristics of the residual treated, forinstance, the type and amount of metal contaminants, and the amount ofnitrogen, sulfur, etc. present. r)The amount of catalyst employed isgenerally from about 0.01 to 5 percent by Weight of the residual treatedwith a preferred amount being about 0.1 to 2 percent.

The hydrogen donor can be contacted with the residual oil to be treatedin any suitable manner. The donor may be added to the oil prior orsubsequent to the addition of the catalyst to the oil. The amount ofhydnoaromatic compound employed is generally at least about 20% of theresidual feed and usually is in a range of from about 50 to 200 percentby Weight of the oil treated. The hydroaromatic material is a liquid atthe conditions of the process and acts as a hydrogen donor as describedabove. When added in admixture with other hydrocarbons the hydroaromaticis usually at least about 40 or 50%, preferably at least about 75% ofthe mixture. During the CDC treatment the temperature is usually in therange of about 700 to 1200 F. with a preferred temperature being about750-850 F. The pressure may vary Widely depending on the particular`feedstock undergoing treatment and the temperature employed, but it isessential that substantial cracking and conversion to lower boiling oilsoccurs. The pressures will generally be elevated and vary from aboutatmospheric to 2000 p.s.i.g. with a preferred range being about 500 to1500 p.s.i.g.

The addition of tree hydrogen in the CDC process is normallyadvantageous as it can increase the liquid product yield and aid in thehydrocracking. Under normal operating conditions it may be desirable toemploy hydrogen, preferably at a partial pressure of at least aboutp.s.i.g. rthere does not appear to 4be any particular benefit in4providing a hydrogen partial pressure in excess of about 1500 psig.,the preferred pressure being from about 300 to 1000 p.s.i.g. Hydrogenconsumption will usually be no more than about 1000 standard cubic feetper barrel of residual treated. This process allows for a conversion ofat least about 20% of the feedstock to a liquid material boiling belowabout 950 F.

Hydrogenation of the residual using free hydrogen and a catalyst may beconducted by contacting the petroleum feed with the catalyst in thepresence of free hydrogen under superatmospheric pressure. Thehydrogenation catalysts generally kno-wn in the art lcan be employed.Calcined solid hyd-rogenation catalysts are preferred and they areusually disposed as a fixed bed o-f macrosized particles, say of about1/a to 1,2" in diameter and about 1/s to 1" or more in length. A movingbed of macrosized catalyst or a luidized bed Iof finely dividedparticles can be used. The catalyst contains catalystically activeamounts of a hydrogenation promoting metal, for instance a heavy metalcomponent such as those of metals having atomic numbers of about 23 to28, the Group VIII catalysts of the iron group, molybdenum, tungsten andcombinations thereof. Frequently the metals are disposed as inorganiccompo- .supported on a solid carrier exemplied by alumina, silica,

etc. Advantageously, the catalyst contains a combination of metals of4the iron group with vanadium or a metal of Group Vla of the periodicchart having atomic numbers from 42 to 74, i.e., molybdenum andtungsten.v A commercial catalyst contains cobalt and molybdenum, eg.,cobalt molybdate, supported on alumina. The amount of catalyticallyactive metal in the supported catalysts is usually about 1 to 30 Weightpercent of the catalyst and preferably about 3. to 20 Weight percent,with there being at least about 1%, preferably at least about 2%, ofeach catalyti'cally active metal when combinations are used.

Catalytic hydrogenation conditions are selected to give the desiredhydrogen consumption and poisoning metals reduction. In general,however, an elevated temperature such as about 600 to 900 F. may beemployed and the pressures are generally superatmospheric usuallyfalling in `the range of about 300 to 3000 p.s.i.g. Free or molecularhydrogen may be provided in the operation and generally in an amount ofabout 50 to 20,000 standard cubic feet per barrel of hydrocarbon toilfeedstock, while the space velocity will lie in the area of about 0.1 tol or more WHSl/v (Weight of hydrocarbon feedstock per hour per weight ofcatalyst). Hydrogen consumption is usually at least about 70-300standard cubic feet of hydrogen per barrel of hydrocarbon oil feed.Where hydrocracidng is desired, such Y as in the treatment of residuals,Vthe hydrogen consumption is often in the range of about 1000 to 2000ermore standard cubic feet per barrel. Residual oils are .often treated at.about 750 to 900 F., at a pressure over about 1000 p.s.i.g., preferablyabout 1500 to 2500 p.s.i.g., and about 100 to 10,000 standard cubicrfeetof hydrogen per barrel.

The treatment of residual material under non-cracking conditions in thepresence of catalytic material and hydrogen or a hydrogen donor isconducted under suitable temperature and pressure conditions so thatthere is not a substantial amount of cracking, ie., less than aboutweight percent of the petroleum feedstock, preferably less than about 5%is cracked. The treating temperature Will usually be in the range ofabout 400 to 700 F., with about 600 to 700 F. being most suitable. Thetotal pressure in the reactor `is usually at least about 100 p.s.i. g.,more often at least about 300 p.s.i.g., and no reason has been seen forgoing above about 3000 p.s.i.g. Preferably the pressure will not beabove about 1500 to 2000 p.s.i.g. It is preferred Vnot to introduce freehydrogen into the system when treating residuals With -a hydrogen donordiluent because the added hydrogen is not consumed efficiently. However,when free hydrogen is introduced into the system the hydrogen partialpressure is about 100-2000 or 3000 p.s.i.g., preferably about G-500lpsig. The length of time of the treatment may vary `Widely so long asconversion of the petroleum feed is limited as noted before. Thetreatment may take from about .1 to 5 hours or more and seems of littlebeneit after 10 hours. The preferred time is about 0.5 to 3 hours and,of course, lower tempera-tures may require longer contact times toobtain a given result.

As in HDC, the hydroaromatic compound can be contacted with thehydrocarbon oil to be treated in any suiti tially hydrogen iodide whichcan be added as such to theV reaction zone or iodine or another hydrogeniodide-producing material maybe added. ln any event, the hydrogen iodideis apparently in equilibrium with elemental iodine in the reaction zonealthough the catalyst may be predominantly hydrogen iodide. rl`hecatalyst can be contacted with the petroleum feedstock in any convenientmanner and the catalyst is essentially'in anhydrous form although it maybe used in solution with alcohol or other solvents. The amount ofcatalystrused can depend upon the reaction conditi-ons and the amount offeedstock demetallization required, but is generally from about 0.1 to20% of the oil to be hydrotreated, preferably about 1 to 10%. Y

After the hydrogenation the hydrogenatcd product material may befractionated usuallyvat atmospheric pressure ,to obtain gaseoushydrocarbons, gasoline, a diluent cut and heavy bottoms comprising theentire oil fraction and material boiling above about it. The gasolinefraction may be hydroformed, thereby increasing itsoctane rating. Thediluent cut is separated from other 'portions of the hydrotreatereffluent and hydrotreated to increase its hydrogento-carbon ratiosuflicient for reuse as a hydrogen donor. Such hydrogenation :may be bythe use of free hydrogen Vin the presenceV of a solid catalyst asdescribed above.

The entire hydrogenated bottoms product described above may be chargedto a vacuum distillation unit, but it is preferable from the standpointof economy in the process of this invention to remove in a distillationat atmospheric or slightly reduced pressure the substantiallymetals-free gas oil fraction boiling essentially in the range of about400 to 700er 900 F., which is amenable to direct catalytic cracking fromthe bottoms which have a boiling point in the range above about 700 to900 F. before vacuum distillation and conduct it directly to thecracking unit. The bottoms are then treated in the vacuum distillationunit to further produce a gas oil fraction boiling primarily in therange -of 700 or 900 F. to 1050 or 1100 Vl?. (corrected to atmosphericpressure) and reduced in metal contaminant content. Thevcontaminatedresidue asphalt fraction fro-m this vacuum distillation generally boilsprimarily above about 1050 or 1100 F. and may be put to low value uses,such as heating oils or road surfacing materials.

Hydrogenation gives a partial reduction in metals content of theresidual feed. The metals remaining in the hydrogenated product wouldaccumulate on the catalyst during the cracking operation and unlesssteps are taken to prevent excess accumulation, excessivedehydrogenation takes place in the cracking, partially undoing the Workperformed in the hydro genating step and severely reducing the yield 4ofgasoline in the cracker eluent. In this invention hydrogenation mayremove only about 10% of the poisoning metal in the residual feed, butpreferably much moreof the poison. Thus the hydrogenated product perhapscontains Vabout 50 to l0 or less Weight percent as much nickel andvanadium as the hydrocarbon charged to the hydrogenating reaction;preferably there is this much reduction in nickel, vanadium or both ofthese metals. `)Frequently the reduction of one or all of nickel,vanadium and iron Will be about to 90 Weight percent. The hydrogenatedproduct contains at least about 2 ppm. nickel and/or about 3 ppm.vanadium, more usually about 25 to 50 ppm. total nickel and vanadium,but rarely more than ptpm. I

Hydrogenation usually does not remove metal contaminants to a point that`is insigniicant in subsequent catalytic cracking. The hydrotreatedproduct also may 'contain some hydrocarbon constituents which areunsuitable for inclusion in a catalytic cracking feedstock. However, bysubjecting the heavy constituents in the hydrotreater eiiuent to vacuumdistillation-a very deep cut of heavy gas oil is obtained Which isreduced in metals and coke formers and therefore suitable as a catalyticcracldng charge to the cracking unit of this invention. The residue fromthe vacuum tower, containing substantial amounts of all the metalcontaminants that were present in the hydrogenated product or fractionthereof, may be removed from the system or recycled back to thehydrogenation unit. Temperature and pressure .conditions are adjusted tosecure the proper distillation of the hydrogenated product or portionthereof. Thus, the

Within the tolerance of the unit for poison.

In theV treatment to take poisoning metals from the vcracking catalyst alarge or small amount of metal can Ibe removed as desired. Thedemetallization treatment generally removes about -to 90% of one or morepoisoning metals from a catalyst portion `which passes through thetreatment. Advantageously a demetallization system is used which removesabout 60 to 90% nickel and 20-40% vanadium from the treated portion ofcatalyst. Preferably at least 50% of the equilibrium nickel content and15% of the equilibrium Vanadium content is removed. The actual time orextent of treating depends on various factors, and is controlled by theoperator according to the situation he facese.g., the extent of metalscontent in the feed, the level of conversion unit tolerance for poison,the sensitivity of the particular catalyst toward a particular phase ofthe demetallization procedure, etc. Also, the thoroughness of treatmentof any lquantum of catalyst in commercial practice is balanced againstthe demetallization rate chosen; that is, the amount of catalyst,

treating rate may be about 5 to 50% of the total catalyst inventory inthe system, per twenty-four hour day of operation although othertreating rates may be used. With a continuously circulating catalyststream, such as 'in the ordinary fluid system a slip-stream of catalyst,

at the equilibrium level of poisoning metals may be removedintermittently or continuously from the regenerator standpipe of thecracking system. The catalyst is subjected to one or more of thedemetallization procedures kdescribed hereinafter and then the catalyst,substantially reduced in contaminating metal content, is returned to thecracking system.

The demetallization of the catalyst will generally include one or moreprocessing steps. Copending patent applications Serial Nos. 758,681,tiled September 3, 1958; 763,833 and 763,834, tiled September 29, 1958;767,794,

tiled October 17, 1958; 842,618, led September 28, 1959; v

849,199, tiled October 28, 1959; 19,313, filed April 1, 1960; 39,810,tiled lune 30, 1960; 47,598, filed August 4, 1960; 53,380, tiledSeptember 1, 1960; 53,623, tiled Septenrber 2, 1960; 54,3868; 54,405 and54,532, filed September 7, 1960; 55,129; 55,160 and 55,184, tiledSeptember 12, 1960; 55,703, filed September 13, 1960; 55,838,

tiled September 14, 1960; 67,518,1iled November 7, 1960;

73,199, tiled December 2, 1960; and 81,256 and 81,257, led January 9,1961; all of which are hereby incorporated by reference, describeprocedures by which vanadium and other poisoningmetals included in asolid oxide hydrocarbon conversion catalyst are removed by dissolving-them from the catalyst or subjecting the catalyst, outside thehydrocarbon conversion system, to elevatedV temperature conditions whichput the metal contaminants into the chloride, sulfate or other volatile,water-dispersible or more available form. A signicant advantage of theseprocesses lies in the -fact that the overall metals removalV operation,even if repeated, does not unduly deleterious'ly affect the activity,selectivity, pore structure Vand other desirable characteristics of thecatalyst. Y

Treatment of the regenerated catalyst withy rnc-lecuialoxygen-containinggas is lemployed to improve the re- This moval of vanadium from thepoisoned catalyst. treatment is described in copending applicationSerial No. 19,313, and is preferably performed at a temperature at leastabout 50 F. higher than the regeneration temperature, that is, theaveragetemperature at which the major portion of carbon is removed fromthe catalyst. The temperature of treatment with molecularoxygen-contain- Vacteristics of the equipment used.

ing gas will generally be in the range of about 1000 to 1800? F. butbelow a temperature Where thecatalyst undergoes any substantialdeleterious chance in its physical or chemical characteristics,preferably a temperature of about 1150 to 1350 F. or even as high as1600 F. The duration of the oxygen treatment and the arnount vofvanadium prepared by the treatment for subsequent removal is dependentupon the temperature and the char- If any significant amount of carbonis present in the catalyst at the start of this high-temperaturetreatment, the essentialv oxygen contact is that continued after carbonremoval, which may vary from the short time necessary to produce anobservable effect in the later treatment, say, a quarter of an hour to atime just long enough not to damage the catalyst. ln lany event, aftercarbon removal, the oxygen treatment of the essentially carbon-freecatalyst is at least long enough to stabilize a substantial amount or"vanadium` im its highest valence state, as evidenced by a significantincrease, say at least about` 10%, preferably at least about in thevanadium removal in subsequent stages of the process. This increase isover and above that which would have been obtained by the other l metalsremov steps without the oxygen treatment. The maximum practical time oftreatment will vary from about 4 -to 24 hours, depending on the type ofequipment used. The oxygen-containing gas used in the treatment containsmolecular Vxygen as the essential active ingredient and there is littlesignificant consumption of oxygen in the treatment. The gas may beoxygen, or a mixture orc oxygen with inert gas, such as air oroxygen-enriched air, containing at least about 1%, preferably at leastabout 10% O2. The partial pressure of oxygen in the treating gas mayrange Widely, for example, from about 0.1 to 30 atmospheres, but usuallythe total gas pressure will not exceed about 25 atmospheres. Thecatalyst may pass directly from the oxygen treatment to a vanadiumremoval treatment especially Where this is the only importantcontaminant, as may be the case when a feed is derived, for example,from Venezuelan crude. Such treatment may be a basic aqueous Wash suchas described in copending patent applications Serial No. 767,794, andSerial No. 39,810. Alternatively vanadium may be re-V moved by achlorination procedure as described in cepending application Serial No.849,199.

Vanadium may be removed from the catalyst after the high te peraturetreatment with molecular oxygen-containing gas by washing ittvtlr abasic aqueous solution. The pH is frequently greater than about 7.5 andpreferably the solution contains ammonium ions which maybe NH;+ ions ororganic-substituted NH.,Jr ions such as methyl ammonium and quaternaryhydrocarbon radical ammoniums. The amount of ammonium ion in itesolution is suicient to give the desired vanadium removal and will oftenbe in the range of about 1 to 25 or more pounds per ton of catalysttreated. The temperature of the Wash solution may vary within Widelimits: room temperature or below, or higher. `Temperatures above 215 F.require pressurized equipment, the cost of which 'does not appear to bejustitied. Very short contact times,

for example, about a minute, are satisfactory, While the time of Washingmay last 2 to 5 hours or longer. After the ammonium wash the catalystslurry can be filtered to give a cake which may be reslurried with Wateror rinsed in other ways, such as, for example, by a water Wash on thelilter, and the rinsing may be repeated, if desired,V several times.

lAlternatively, after the highv temperature treatment withoxygen-containinggas, treatment ol a metals contaminated catalyst with achlorinating agent at moderately elevated temperature up to about 1000"E'. is of value in removing vanadium contaminants from the catalyst asvolatile chlorides. This treatment is described in cepend- Vingapplication Serial No. 849,199. The chlorination takes place-at atemperature of at least about 300 F.,

preferably about 550 to 650 F. with optimum results usually beingobtained near o F. The chlorinating agent is essentially anhydrous, thatis, if changed to the liquid state no separate aqueous phase would beobserved in the reagent.

The chlorinating Las vapor which contains chlorine or sometimes rably incombination with carbon or suiur. Such reagents include molecular ccrine but preferably are i res chlorine with, for example, a chlorinesubstituted light hydrocarbon, such as carbon tetrachloride, which maybe used as such or formed in situ by the use or", for example, avaporous mixture of chlorine gas with lo' molecular weight hydrocarbonssuch as othane, n-pen:ane, etc. About li0 percent active chlorrnatingagent based on the weight of the catalyst is generally used. carbon orsulfur co.-.- oounrl promoter is generally used in the amount ot aboutor l0 percent or more, preferably about 2 3 percent, ased or. the Weigdtof tie catalyst for good metals renoval; however, even ir less t? thisamount is used,

considerable imp ovemcnt in metals conversion is obtained over thatwhich is `possible at the same temperature using chlorine alone.chlorine and promoter may be supplied individually' or as a mixture to apoisoi ed catalyst. Such a mixture may contain about 0.1 to 50 p scluorine per part of promoter, preferably about l-l0 parts per part ofpromoter. A chlorinating gas comprising about l30 weight percentchlorine, based on the catalyst, together with one percent or more SgClggives good results. lareferably, such a gas provides 1-l0 percent Clgabout 1.5 percent SgClg, basen on the catalyst. A saturated mixture ofCCL; and Clz or HC1 can be made by bubbling chlorine or hydrogenchloride at room temperr. through a vessel containing such a mixturegenerally contains about l part 0 parts C12 or HCl. Conveniently, apressure 0400 or more p.s.i preferably about 0.15 may be maintained inchlorination. The chlorinay taire about 5 to l2() minutes, more usuallyabout u minutes, but shorter or longer reaction periods `tay be possibleor needed, for instance, depending on the linear velocity of thecblorinating and purging vapors.

"file demetallization procedure employed in this invention be directedtoward nickel removal from the catalyst, generally in conjunction withvanadium removal. Nickel removal may be accomplished by dissolvingnickel compounds directly the catalyst and/or by convertthe nickelcompounds to volatile materials and/or materials soluble or dispersiblein an aqueous medium, eg., Water or dilute acid. rhe Water-dispersibleorrn may be one which decomposes in water to produce Watersoluble.products. The removal procedure for the converted metal :nay be based onthe form to which the metal is converted. The mechanism of the Washingsteps may be one of simultaneous conversion of nickel and/or vanadium tosalt or other dispersible form and removal by tbe aqueous Wash; however,this invention is not to be limited by such a theory.

Conversion of some of the metal poisons especially nickel, to awater-dispersible form is described in copending ap'ilication Serial No.758,681 and may be accompli ned, for instance, by subjecting thecatalyst to a suliating gas, that is S132, S03 or a mixture of SO2 andO5, at an elevated temperature. Sulfur oxide Contact is usuallypzrformed at a temperature of about S00 to l200 F. and frequently it isadvantageous to include some free oxygen in the treating gas. Anotherprocedure, described in copending application Serial No. 763,834 andSerial No. $42,618, includes sulding the catalyst and performing anoxidation process, after which metal contaminants in Water-dispersibleform, preferably prior to an wash may be removed from the catalyst by anaqueous medium.

The suliiding sep can be performed by contacting the `zoistuied catalystwith elemental sulfur vapors, or more A tion.

conveniently by contacting the poisoned catalyst with a volatilesulfide, such as HES, CS2 or a niercaptan. The Contact with thesulfurcontaining vapor can be performed at an elevated temperaturegenerally in the range oi about 500 to 1500" F., preferably about 800 to1300o F. Other treating conditions can include a sulfur-containing4vapor partial pressure of about 0.1 to 30 atmospheres or more,preferably about 0.5 to 25 atmospheres. Hydrogen sulfide is thepreferred suliiding agent. Pressures below atmospheric can be obtainedeither by using a partial vacuum or by diluting the vapor with gas suchas nitrogen or hydrogen. The time of contact may vary ou the basis ofthe temperature and pressure chosen and other factors such as the.amount of metal to be removed. The sulding may run for, say up to about20 hours or more depending on these conditions and the severity of thepoisoninc. Temperatures of about 900 to 1200 F. and pressuresapproximating l atmosphere or lless seem near optimum for sulding andthis treatment often con tinues for at least 1 or 2 hours but the time,or course, can depend upon the manner of contacting the catalyst andsuliidiug agent and the nature of the treating system, e.g., batch orcontinuous, as Well as the rate of diiiusion Within the catalyst matrix.The sulding step performs the function not only of supplying asulfur-containing metal compound which may be easily converted to aWater-dispersible form but also appears to concentrate some metalpoisons, especially nickel, at the surface of the catalyst particle.

Oxidation after sulding may be performed by a gaseous oxidizing agent toprovide metal poisons in a dispersible form. Gase-ous oxygen, ormixtures of gaseous oxygen with inert gases such as nitrogen, may bebrought into contact with the sulfided catalyst at au oxygen partialprcssure of about 0.2 atmosphere and upward, temperatures upward of roomtemperature and usually not above about 1300" F., and times dependent ontemperature and oxygen partial pressure. Gaseous oxidation is bestcarried out near 900 F., about one atmosphere O2 and at very briefcontact times.

The metal su'lde may be rendered Water-dispersiblc by a liquid aqueousoxidizing agent such as a dilute hydroxen peroxide or hypochlorous acidwater solution, as described in copending application Serial No.842,618. The inclusion in the liquid aqueous oxidizing solution ofsulfuric acid or nitric acid has been found lgreatly to reduce theconsumption of peroxide. In addition the inclusion of nitric acid in theoxidizing solution provides for increased vanadium removal. Usefulproportions of acid to peroxide to catalyst generally include about 2 to25 pounds acid (on a 100% basis) to about 1 to 30 pounds or more H2O2(also on a 100% basis) in a very dilute aqueous solution, to about oneton of catalyst. A 30% H2O2 solution in Water seems to be anadvantageous raw material for preparing the aqueous oxidizing solu-Sodium peroxide or potassium peroxide may be used -in place of hydrogenperoxide and in such circumstances, enough extra sulfuric or nitric acidmay be used to provide one mole of sulfate or two ruols of nitrate foreach two mois of sodium or potassium.

Another highly advantageous oxidizing medium is an aerated dilute nitricacid solution in water. Such a solution may be provided by continuouslybubbling air into a slurry of the catalyst in very dilute nitric acid.Other oxygen-containing gases may be substituted for air. Varying oxygenpartial pressure in the range of about 0.2 to 1.0 atmosphere appears tohave no eiiect in time required for oxidation, which is generally atleast about 7 to 8 minutes. The oxidizing slurry may contain about 20%solids and provide about iive to ten pounds of nitric acid per ton ofcatalyst. Studies have shown a greater concentration of HNO3 to be of nosignificant advantage. Other oxidizing agents, such as chromic acidWhere a small residual CrgOg content in the catalyst is not signincant,and similar aqueous oxidizing solutions such as wav surface.

V aieaao `i553 ter solutions of manganates and permanganates, chlorites,chlorates and perchlorates, bromites, bromates and perbromates, iodites,iodates and periodates, are also useful. Bromine or iodine water, oraerated, ozonated or oxygenated water, with or without acid, also willprovide a dispersible form. The conditions of oxidation can be selectedas desired. The temperature can conveniently range up to about 220 F.with temperatures of above about 150 F. being preferred. Temperaturesabove about 220 F. necessitate the use of superatmospheric pressures andno need for such has been found.

After provision of nickel sulfide in a dispersible form, the catalyst iswashed with an Yaqueous medium to remove the metal poisons. This aqueousmedium, for best removal of nickel is generally somewhat acidic. Theaqueous medium cancontain extraneous ingredients in trace amounts, solong as the medium is essentially water and the extraneous ingredientsdo not interfere with demetallization or adversely affect the propertiesof the catalyst. Ambient temperatures can be used in the wash buttemperatures of about 150 F. to the boiling point of water are sometimeshelpful Pressures above atmospheric may be used but the results usuallydo not justify the additional equipment. Where an aqueous oxidizingsolution is used, the solution may perform part or all f the metalcompound removal simultaneously with the oxidation'. In Order to avoidundue solution of alumina from a chlorinated catalyst, contact time inthis stage is preferably held to about 3 to 5 minutes which issufficient for nickel removal. Also, since a slightly acidic solution isdesirable for nickel removal, this wash preferably takes place beforethe ammonium wash.

Alternative to the removal ofV poisoning metals by procedures involvingcontact of the suliided or sulfated catalyst with aqueous media, nickelpoison may be removed through conversion of the nickel sulfide to thevolatile nickel carbonyl by treatment with carbon monoxide, as describedin copending application Serial No. 47,598. In such a procedure thecatalyst is treated with hydrogen at an elevated temperature duringwhichV nickel contaminant is reduced to the elemental state, thentreated, preferably under elevated pressure and at a lower temperaturewith carbon monoxide, during which nickel carbonyl is formed and flushedolf the catalyst Hydrogenation takes place ata temperature of about 800to 1600 F., at a pressure from atmospheric or less up to about 1000p.s.i.g. with a vapor containing 1 0 to 100% hydrogen. Preferredconditions are a pressure up to about p.s.i.g. and a temperature ofabout 1100 to l300 F. and a hydrogen content greater than about 80 molpercent. AThe hydrogenation is continued until surface accumulations ofpoisoning metals, particularly nickel, are substantially reduced to theelemental state. Carbonylation takes place at a temperaturesubstantially lower than the hydrogenation,

from about ambient temperature to 300 F. maximumY and at a pressure upto about 2000 p.s.i.g with a gas i containing `about 50-100 mol percentCO. Preferred conditions include greater than about 90 mol percent CO, apressure of up to about 800 p.s.i.g. and a temperature of about -100180F. The CO treatment serves generally both to convert the elementalmetals, especially nickel Y to volatile carbonyl and to remove thecarbonyl.

After the ammonium wash, or after the final treatment which may be usedin the catalyst demetallizationl procedure, the catalyst is conductedback to the crackmg system. Where a small amount of the catalystinventory is demetallized, the catalyst may be returned to lest thewater put out the lire or unduly lower the temperature in theregenerator, it may be desirable first Vto dry a wet catalyst filtercake or filter cake slurry at say about 250 to 450 IE. andV also, priorto reusing the f catalyst in the cracking operation it can be calcined,`say at temperatures usually in the range of about 700 to 1300 F.Prolonged calcination of the catalyst at above about 1l00 F. maysometimes be disadvantageous. Calcination removes free water, if any ispresent, and perhaps some but not all of the combined water, and leavesthe catalyst in an active state without undue sintering of its surface.Inert gases such as nitrogen frequently may be employed after contactwith reactive vapors to remove any ofthese ,vapors entrained in thecatalyst or to purge the catalyst of reaction products.

The demetallization procedure of thisrinvention has been found to behighly successful when used in conjunction with tluidized catalyticcracking'systems to control theamount of metal poisons on the catalyst.When such catalysts are processed, a fluidized solids technique isrecommended for these vapor contact demetallization procedures as a wayto shorten the time requirements. Any given step in the demetallizationtreatment is Vusually continued for a time sufficient to effect asubstantial conversion or removal of poisoning metal and ultimatelyresults in a substantial increase in metals removal compared with thatwhich would have been removed if the particular step had not beenperformed. After the available catalytically active poisoning metal hasbeen removed, in any removal procedure, further reaction time may haverelatively little effect on the catalytic activity of the depoisonedcatalyst, although further metals coutent may be removed by repeated orother treatments.

This invention will be better understood by reference to theaccompanying drawings which shows the schematic of a representativeprocessing system but is not to be construed as limiting.

A residual lfeed contaminated with poisoningmetals, for example, avacuum asphalt, is fed by line 10 to hydrogenation unit 12. Thestructure of the unit will, of course, depend upon the hydrogenationprocess; for instance in 'the HDDC method, the unit l2 may consist of acoil or a coil and drum. ln this method the feed is heated in the coilto about 700 to 1200 F. and then contacted in the drum with a donordiluent,`such as tetralin. metals removed will, of course, depend uponthe severity of the hydrogenation and the characteristics of theoriginal feed. In the alternative hydrogenation procedures,hydrogenation may be conducted under other conditions previouslydescribed. The hydrogen donor enters hydrogenation unit 12 by line 14.Line 15 is provided for the introduction of free hydrogen, i-f any isemployed. Suitable pumps, not shown, may be provided to mix the residualfeed with the hydrogen donor. The conditions in the unit, l2, areadjusted for the results required. The total products are passed fromthe unit 12 to fractionator 16 by line 18. AFixed gases are removed byline 20 and gasoline and lighter components having an end boiling pointof about 430 F. are removed by line 22 to storage or further treatmentsuch as hydroforming. The hydrogen donor diluent may be removed from thefractionator by line 24 and hydrogenated by means, not shown, beforebeing reintroduced intoV the system by line 14 withfresh hydrogen Adonordiluent, if needed, from line 25. Gas oils having .a boiling pointwithin the range of about 400 to L700 or 800 F. may be removed by line2,6, and carried directly to the catalytic i cracker 28 by lines 30 and32 or it may be removed from fractionator 16 along with'the bottomsfraction, which usually has a boiling point essentially above about 700or 800 F. by line 34 and conducted to vacuum tower 36, where thecombined gas oil and bottoms fractions or the'bottoms fraction alonemayV be distilled at temperatures of about 750 F. and pressures of about25 mm. Hg. Steam may be introduced by line 38. The residue, containingessentially heavier asphaltic constituents boiling above about 1050 or1100 F., is removed from the vacuum tower by line 40 and may beconducted back to the hydrotreater by lines 40 and 10 for The extent ofconversion and the amount of further processing or may be Withdrawn fromthe system by line 41. Vacuum gas oil containing cracking components ofreduced metal content, is removed from the vacuum tower by line 42 andconveyed to the cracker 28 by line 32. The cracking feed may be furtherdiluted with low metals content conventional cracking stock from anoutside source by lines 44 and 46. The cracker eluent leaves by line 50and is brought to fractionator 52, where components of the etiluent arewithdrawn by line 54 for fixed gases, line 56 for gasoline, line 58 forgas oil components and line 60 for materials higher boiling than gasoil. The latter components may be Withdrawn from the system by line 62or may be recycled by lines 64, 66, 40 and 10 to the hydrogenatiou zone12, for further processing. The gas oil fraction is preferably recycledto the catalytic cracking step by lines S, 67, 48, 46 and 32 since it issubstantially poison-free, or it may be Withdrawn for use as distillatefuel by line 68.

Contaminated catalyst is continuously removed from the cracker 28 byline 69 which conducts it to the regenerator 70. The regenerator isprovided with the exit 72 for exhaust gases and with line 74% for theremoval of regenerated catalyst and return to the cracker 28. A smallslip stream of catalyst may be removed from line 74 for demetallizationby line 76 and conveyed to demetallization unit 78. The demetallizationunit 78 may comprise a system Which includes apparatus (not shown), forexample, for sulding, chlorinating, washing and ltering the catalyst.Alternatively, instead of chlorinating the sultided catalyst, means foroxidizing the sulded catalyst may be substituted. The demetallizationprocedure used will, of course, depend upon the metals present. When,for instance, the feed to be treated contains predominant amounts ofvanadium, the demetallization treatment will be geared mainly toward theremoval of vanadium such as for example, dernetallization apparatus forhigh temperature treatment with molecular oxygen containing gas asdisclosed in S.N. 19,313 followed by a basic wash as disclosed inapplication S.N. 39,810. After a substantial portion of the metalcontaminants are removed from the cracking catalyst, the catalyst isreturned to the cracking system by line 80.

The treating process of the present invention may be exemplified by thefollowing:

A North Texas reduced crude having an API gravity of about 22, a carboncontent of about 5.3 Weight percent, having an initial boiling pointabove about 650 F. and containing about 25 p.p.m. nickel and 60 p.p.m.of vanadium is mixed with a hydrogen donor diluent comprising ahydrogenated catalytic cycle oil in a l to l donor diluent-to-oil ratioand thermally cracked. The thermal cracking is operated at a temperatureof about 800 to 900 F. and under a pressure ot about 100 p.s.i.g. Thefeed rate is controlled so that the blend of reduced crude and donordiluent is held for about 30 minutes at about 820 F. About 40% of the1050 E+ components of the crude are converted to gas oil and lowerboiling products. The total products are conducted to a fractionatorWhere C4* gases and gasoline having an end boiling point of about 430 F.are taken off at atmospheric pressure. The combined gas oil and bottomsfraction having a metals content of about 12 p.p.m. nickel and 20 p.p.m.vanadium is distilled at a temperature of about 750 F. and a pressure ofabout 25 m.m. Hg. The residue, containing material not suitable forcatalytic cracking, is conveyed back to the hydrogenation zone forfurther processing. The overhead fraction, amounting to about 85% of thefeed to the vacuum still, containing materials boiling above 400 F. andcontaining 2 p.p.m. nickel and 6 p.p.m. vanadium is diluted With arecycle gas oil from the cracker fractionator. The feed to the crackingunit contains about 1.5 p.p.m. Ni() and 4.5 p.p.m. V205. In thecatalytic cracker the feed contacts a synthetic gel silica-aluminacatalyst, having an A1203 content of about 25%, at a temperature ofabout 950 to 975 F. and a pressure of about 5 p.s.i.g. The crackedproducts are introduced to a fractionator where a yield of gasoline andother components are removed. The gas oil fraction is recycled to thecracker. A portion of the silica-alumina catalyst is continuouslyremoved from the cracking reactor and brought to a regenerator. Averageresidence time in the regenerator is about 5 minutes at a temperature ofabout 1100 F. before catalyst return to the reactor at a carbon level ofless than about 0.5%.

About 10% per day of the cracking catalyst inventory poisoned to ametals level of about 305 p.p.m. NiO and 1820 p.p.m. V205 is sent as aside stream from the regenerator to demetallization. In thedemetallization process the catalyst is held in air for about an hour atabout 1300 F. and then sent to a suliiding zone Where it is liuidizedwith H28 gas at a temperature of about 1175 F. for about 1 hour. Dilutenitric acid is brought in contact With the suliided catalyst and theslurry is aerated for about 10 minutes at a temperature of 200 F. toconvert nickel poisons to dispersible form and remove them. The catalystis then washed with an ammonium hydroxide solution having a pH of about8 to l1, removing the available vanadium. The catalyst, substantiallyreduced in nickel and vanadium content is filtered from the wash slurry,dried at about 350 F. and returned to the regenerator. The treatedcatalyst analyzes a metals content of 120 p.p.m. nickel and 1330 p.pm.vanadium.

In another example, an asphaltic residual oil from the vacuumdistillation of the bottoms from an atmospheric distillation ofpetroleum crude oil is contacted with tetralin in a ratio of diluent tofeed of 1 to l, at about 800'J F. and a pressure range of about 11,60 to1260 p.s.i.g. and a hydrogen partial pressure of 500 p.s.i.g. in thepresence of a catalyst comprising about 2% l2. The residual oil has ametals content of 82 p.p.m. Ni() and 244 p.p.m. V205. 30% of the feed isconverted into products boiling below about 950 F. The total productsare conducted to a fractionator operating under slightly reducedpressure, Where about 7% of gas and gasoline are removed, and about 13%of gas oil is recovered and D sent to the cracking unit. The bottomsfraction, boiling essentially above about 800 F. and having a metalscontent of 40 p.p.m. nickel and Sl p.p.m. vanadium is distilled at atemperature of 750 F. under a pressure of about 25 mm. Hg resulting in ayield of 30 percent by weight vacuum gas oil having a metals content ofl p.p.m. nickel and 2 p.p.m. vanadium. The vacuum gas oil is dilutedwith the gas oil removed in the first fractionator, a recycle gas oiland a conventional cracking feed to give a cracking feed containing 0.75p.p.m. nickel and 1.5 p.p.m. vanadium.

In the catalytic cracker the feed contacts a syntheticgel silica-aluminacatalyst, having an A1203 content ot about 25 at a temperature of about950 to 975 F. and a pressure of about 5 p.s.i.g. The cracked productsare introduced to a fractionator Where a yield of gasoline and othercomponents are removed. The gas oil is recycled to the cracker forfurther processing. A portion ofV the silica-alumina catalyst iscontinuously removed from the cracking reactor and brought to aregenerator. Average residence time in the regenerator is about 5minutes at a tem erature of about l F. before returning to the reactorata carbon .level of less than about 0.4%.

About 5 per day of the cracking catalyst inventory poisoned to a metalslevel of about 280 p.p.m. NiO and 1025 p.p.m. vanadium is sentas a sidestream from the regenerator to demetallization. In thev demetallizationprocess the catalyst is held in air for about an hour at about l300 F.and then sent to -a suliiding zone Where it is iluidized with H28 gas ata temperature of about tion having a pH of about 8 to ll.

Vreturned to the 'regeneraton analyzed and `shows a metals content of 90ppnrn.'nickel and '665 ppqm. vanadium..

Thus, this Yinvention providesrforovercoming Apoison- 'irig eects by abalanced process whichV includes hydro- ,Y

- fai- 1175 F. for about 1 hour. The catalyst is then purged with iluegas ata temperature of about 575 F. and chlorinated Vin a chlorinationyzone with an equimolar mixture of C12 andCCl4 at about 600o After about1 hour no trace of vanadiumrchloride can be found in the chlorinationeliiuent and the catalyst is quickly washed with water. A pH of about 2is imparted to this wash medium by chlorine entrained in the catalystand the wash serves to remove nickel chloride. The demetallizationprocedure removes about 60% ot the nickel and about 25% of the vanadiumon the catalyst.

A third Vrun is conducted with a West Texas asphalt having an initialboiling point of about 950 F., a specific gravity of 1.002, a Conradsoncarbon content of about 20.5 and a metalslevel of 76'p.p.m. nickel oxideand 110 ppm. vanadium oxide. The feed is diluted with a substantiallymetals-free cycle oil in a ratio of 1 to land sent to a hydrotrea-tingoperation as amixture of liquid and vapors. The hydrotreater' contains afixed bed of cobalt-molybdena-alumina catalyst analyzing about 2.5% Coand 9% M003. In the hydrotreater Vthe conditions maintainedl are about725 F., about-500 p.s.i.g., about 1 WHSV and 300 standard cubic feet ofhydrogen per barrel of feed, about 250 s.c.r". of. which are consumed.Cracking is held to a minimum and onlyabout of the asphalt feed isconverted to products boiling below i about 950 F. VHydrogen consumptionis about 437 S.c.f./bbl. of feed. vThe hydrogenated product, analyzing Yfraction containing a metal contaminant selected from the groupconsisting of about 1-5 ppm. nickel and about 1-10 ppm. vanadium, saidnickel and vanadium being calculated asktheV metal oxides, and anundistilled residue boiling primarily above about 1100 F. and containingasphaltic material, subjecting tocatalytic cracking a hydrocarbonfeedstock containing said gas oil fraction in an amount to provide atleast about 0.5 p.p.m. of said contaminating metal, in the presence of asolid cracking catalyst under cracking conditions to produce gasoline,removing metal contaminated catalyst from the cracking system, theremoved catalyst containing at least about 50 p.p.rn. of a metalcontaminant selected from the group vconsisting of nickel and vanadium,demetallizing removed a hydrogen donorzdiluent iny proportions of about0.1 to

about 5.5 ppm. nickeloxide and 10.2 ppm. vanadium i' oxide, isvacuum'distilled at a temperature of 750 F.

and pressure or" mm. Hg and the resulting gras oil boil-y ing betweenabout 400 to 1100 F. and amounting to about 35% of the hydrogenatedproduct, analyzes 1.5 p.p.m. nickel oxide and 3.0 ppm. vanadiumV oxide.The vacuurngas oil is catalytically cracked at a temperature ot about950l`.,rat 10 psig. pressure in the presence 13% A1203; The crackedproducts are introduced to a fr'actionator Where a 60% yield ofgasolineand other low boiling components are removed. The products boiling aboveabout 950 F. are recycled to the hydrogenator forV further processing. Aportion ofthe silicaalumina catalyst is continuously removed from thecracking reactor' yand brought to a regenerator. Average residence timein the regenerator is about 5 minutes at aternperature of about 1100JF., before returning V'to `.the reactor at a carbon level of about 0.5%.A

About 10% of the cracking catalyst inventory, poisoned to a metals levelor" about 225 ppm. nickeland 910 ppm. vanadium, is each day sent as aside stream fromV the regenerator to demetallization. process thecatalyst is held in air for about an hour 'at about 1300 F. then inHzSfor about an hour at 1150 F., and in a C12/CS2 vapor mixture at 600 F.and then sent to a basic wash with an ammonium hydroxide solu- Tljrecatalyst is iiitered from the'washV slurry, dried at about y350 F. andyi TheA treated catalyst is genation, distillation and craclcng catalystdemetallization. it is claimed: i 1. A process fortreatirig' a residualhydrocarbon oil boiling .above the ygasoline range and containing atleast about ,5 p.p.m. of a metal contaminantselected from Vthe Vgroupconsisting of nickel and vanadium in an amount suiiicienttoV causedeterioration in selectivity ot a cracking catalyst, comprising thesteps of treating said hydrocanbonoil in a hydrogenation zone to reduce'by aboutvv In the demetallizationk 10 volumesY of diluent per volume ofresidual hydrocarbon oil, subjecting the mixture. to temperatures in theVrange of .about 700-l200 F. and pressures from about atmospheric toabout 2000 psig., and a hydrogenated is hydrogen bromide. Y

' of a synthetic-gel silica-alumina catalyst containing about l y. 6.The process of claim 1 wherein hydrogenation is carried out bycontacting the residual hydrocarbon oil at temperatures of about 600 to900 F. and pressures of about 300 to 3000 p.s.i.g. with molecularhydrogen and Va hydrogenation catalyst.

7. The process of claim 6 wherein the hydrogenation` catalyst is a solidoxide offa ,hydrogenation promoting met-al on asolid carrier. Y

8. The process of claim 1V wherein cracking catalyst de- -metallizationincludes contact ot` thecatalyst with a vapor reactive with a metalcontaminant.

9. The process of claimrl wherein the solid cracking 1 synthetic gelsilica cracking catalyst lunder cracking cond-i4 10 -190% the content ofsaid contaminantl metal, subiecting'a thus hydrogenated fraction boilingessentially above A about 400 to` vacuum-distillationto' separate a gasoil tionsV to produce gasoline, removing from the cracking systemcatalyst contaminated with said selected cont-ami-` nant, demetallizingthe catalyst, returning catalyst to said l cracking system, andrecovering gasoline product from saidcracking; saidfeedstockcontaining,the vacuum gas oil traction obtainedwby treating theresidual hydrocarbon Yoilboiling aboveV the gasoline rangeV andcontainingl vrnetal contaminantv in a'fhydrogenation zoneYto reducepartially the content offsaidgcontaminant metal by at least about 10%Vandsubjecting a hydrogenated traction boiling essentiallyV aboveV about400 F. to vacuum distil-` lation at artemperature'of about 7 00-775,F.and'a pressure ofy about 10-30 mm. Hg.

L12. A process Y l boiling,v above the gasoline range and containingabout 10 to-500 ppm; nickel and aboutAlO to 1000 ppm. va#

nadium and in an amount sufficient to cause deterioration in selectivityof a synthetic lgel,silica-based cracking catafor treating a residualhydrocarbon oil.V n

oil ina hydrogenation zone to reduce by about 50 to 90% the content ofsaid contaminant metals, subjecting a thus hydrogenated fraction boilingessentially above about 400 F. to vacuum distillation to separate a gasoil fraction containing about 1-6 p pm. nickel and about l-lO p pm.vanadium, said nickel and Vanadium being calculated as the metal oxides,and an undistilled residue boiling primarily above about l100 F. andcontaining asphaltic material, subjecting to catalytic cracking ahydrocarbon feedstock containing said gas oil fraction in an amount toprovide about 0.5 to less than about 4 p.p.m. nickel and about 0.5 toless than about 8 p,p.m. vanadium, in the presence of a solid, syntheticgel, silica based cracking catalyst under cracking conditions to producegasoline, removing metal contaminated catalyst from the cracking system,the removed catalyst containing at least about 200 p.p.m. of nickel andat least about 500 p.p.m. vanadium, demetallizing removed catalyst toWithdraw at least 50% of the nickel and at least 15% of the vanadium,returning resulting demetallized catalyst to said cracking system andrecovering the products from said cracking.

13. The method of claim 12 wherein the catalyst is silica-alumina.

14. The method of claim 13 in which catalyst from the catalytic crackingis regenerated to remove carbon and demetallized by contact with anoxygen-containing gas at a temperature of about 1000 to 1800u F. toenhance subsequent vanadium removal, sulided by contact with a suldingagent at a temperature ot about 500 to 1500 F. to enhance subsequentnickel removal, chlorinating poisoning metal containing component on thecatalyst by Contact with an essentially anhydrous chlorinating agent ata temperature of about 300 to 1000" F., and contacting the catalyst witha liquid, essentially aqueous medium to remove soluble poisoning metalcomponents from the catalyst.

References Cited in the le of this patent UNITED STATES PATENTS2,758,097 Doherty et al. Aug. 7, 1956 2,768,121 Denton et al Oct. 23,1956 2,928,784 Goldsmith Mar. 15, 196'() 2,944,013 Holden July 5, 19603,008,897 Burk et al Nov. 14, 1961 3,028,331 Tucker Apr. 3, 1962 UNITEDSTATES PATENT OFFICE CERTIFICATE 0F CORRECTION Patent No. 3,162,596December ZZ, 1964 Arvin D. Anderson et al.

It is hereby certified that error appears in the above numbered patentrequiring correction and that the said Letters Patent. should read ascorrected below.

Column 13, line 68, for "application" read applications column :16, line3l, for "drawings" read drawing column 20, line 16, for "resulitng" readresulting line 47, for "calatyst" read -4 catalyst columnl 21, lineSigned and sealed this 29th day. of June 1965.,

(SEAL) Attest:

ERNEST W. SWIDER EDWARD J. BRENNER Attesting Officer Commissioner ofPatents

1. A PROCESS FOR TREATING A RESIDUAL HYDROCARBON OIL BOILING ABOVE THEGASOLINE RANGE AND CONTAINING AT LEAST ABOUT 5 P.P.M. OF A METALCONTAMINANT SELECTED FROM THE GROUP CONSISTING OF NICKEL AND VANADIUMINAN AMOUNT SUFFICIENT TO CAUSE DETERIORATION IN SELECTIVITY OF ACRACKING CATALYST, COMPRISING THE STEPS OF TREATING SAID HYDROCARBON OILIN A HYDROGENATION ZONE TO REDUCE BY ABOUT 10-90% THE CONTENT OF SAIDCONTAMINANT METAL, SUBJECTING A THUS HYDROGENATED FRACTION BOILINGESSENTIALLY ABOVE ABOUT 400* F. TO VACUUM DISTILLATION TO SEPARATE A GASOIL FRACTION CONTAINING A METAL CONTTAMINANT SELECTED FROM THE GROUPCONSISTING OF ABOUT 1-5 P.P.M. NICKEL AND ABOUT 1-10 P.P.M. VANADIUM,SAID NICKEL AND VANADIUM BEING CALCULATED AS THE METAL OXIDES, AND ANUNDISTILLED RESIDUE BOILING PRIMARILY ABOVE ABOUT 1100*F. AND CONTAININGASPHALTIC MATERIAL, SUBJECTING TO CATALYTIC CRACKING A HYDROCARBONFEEDSTOCK CONTAINING SAID GAS OIL FRACTION IN AN AMOUNT TO PROVIDE ATLEAST ABOUT 0.5 P.P.M. OF SAID CONTAMINATING METAL, IN THE PRESENCE OF ASOLID CRACKING CATALYST UNDER CRACKING CONDITIONS TO PRODUCE GASOLINE,REMOVING METAL CONTAMINATED CATAYST FROM THE CRACKING SYSTEM, THEREMOVED CATALYST CONTAINING AT LEAST ABOUT 50 P.P.M. OF A METALCONTAMINANT SELECTED FROM THE GROUP CONSISTING OF NICKEL AND VANADIUM,DEMETALLIZING REMOVED CATALYST TO WITHDRAW ABOUT 10 TO 90% OF SAID METALCONTAMINANT, RETURNING RESULING DEMERALLIZED CATALYST TO SAID CRACKINGSYSTEM AND RECOVERING THE PRODUCTS FROM SAID CRACKING.